Process for the production of olefins

ABSTRACT

Process for the production of an olefin comprising partially combusting in a reaction zone a mixture of a hydrocarbon and an oxygen-containing gas in the presence of a catalyst which is capable of supporting combustion beyond the fuel rich limit of flammability to produce the olefin. The superficial feed velocity of the mixture is at least 250 cm s −1  at standard temperature and operating pressure, with the proviso that where the catalyst is an unsupported catalyst, the superficial feed velocity of the mixture is at least 300 cm s −1  at standard temperature and operating pressure. The process is carried out at a pressure of at least 1.3 bara and the reaction zone is not externally heated.

The present invention relates to a process for the production of olefinsfrom a hydrocarbon feed and, in particular, to a process for theproduction of olefins by the partial combustion of a hydrocarbon feed.

BACKGROUND OF THE INVENTION

Olefins such as ethylene and propylene may be produced by the catalyticdehydrogenation of a hydrocarbon feed or the cracking of a hydrocarbonfeed. The term “cracking” will be used throughout this specification toembrace both of these chemical reactions.

The cracking of hydrocarbons is an endothermic process. Accordingly,heat has to be consumed for the reaction to occur. Auto-thermal crackingis a known process for the production of olefins from a reactant mixturecomprising a hydrocarbon feed and an oxygen-containing gas. An exampleof an auto-thermal cracking process is described in EP-A-0 332 289.

In an auto-thermal cracking process, the heat required for cracking isgenerated by combusting a portion of the original hydrocarbon feed. Thisis achieved by passing a mixture of a hydrocarbon feed and anoxygen-containing gas over catalyst capable of supporting combustionbeyond the fuel rich limit of flammability. The hydrocarbon feed ispartially combusted, and the heat produced by the combustion reaction isused to drive the cracking of the remainder of the feed into olefins.Optionally, a hydrogen co-feed is also burned, and the heat produced bythis combustion reaction is also used to drive the cracking of thehydrocarbon.

In an auto-thermal cracking process, the time for which the reactionmixture (hydrocarbon and an oxygen-containing gas) is in contact withthe catalyst (the contact time) is believed to have an impact on theolefin yield of the overall process. Olefin yield is determined by theselectivity of the process towards olefins and the extent of hydrocarbonconversion. For high olefin yields, high selectivity and high conversionare desirable. In general, the conversion of hydrocarbon increases asthe contact time increases. Without wishing to be bound by any theory,it is believed that this is because there is more time available for thehydrocarbon to react. However, increasing the contact time tends to havea detrimental effect on the selectivity to olefin, as there is more timefor the olefin produced to take part in further (undesirable) reactions.

An indication of contact time can be obtained by measuring the linearvelocity of the feed gases upstream from the catalyst at standardtemperature (273 Kelvin) and the operating pressure of the process. Thismeasurement, known as the superficial feed velocity, is measured incentimetres per second (cm s⁻¹). The higher the superficial feedvelocity, the shorter the contact time of the feed for a given catalystquantity and aspect ratio.

Conventional understanding thus indicates that if high superficial feedvelocities are employed in an auto-thermal cracking process thehydrocarbon feed conversion and olefin yield would be significantlyreduced. Indeed, it would be expected that conversion and olefin yieldwould be reduced to such an extent that any potential benefitsassociated with operation at high superficial feed velocities would benegated.

This teaching has been exemplified by prior art catalytic oxidativedehydrogenation processes. Prior art catalytic oxidative dehydrogenationprocesses have been operated at superficial feed velocities of up to 265cm s⁻¹, but, more typically, such processes are operated at superficialfeed velocities of less than 180 cm s⁻¹.

U.S. Pat. No. 5,639,929 discloses an oxidative dehydrogenation processusing a fluidised bed catalyst of Pt, Rh, Ni or Pt—Au supported onα-Al₂O₃ or ZrO₂ and total feed flow rates of 0.5 to 2.0 SLPM (standardlitres per minute) corresponding to superficial feed velocities of ˜1 to˜4.1 cm s⁻¹ at standard temperature and operating pressure.

U.S. Pat. No. 5,905,180 discloses a catalytic oxidative dehydrogenationprocess wherein the total feed flow rates “ranged from 5 SLPM”,corresponding to a superficial feed velocity of ˜24 cm s⁻¹ at standardtemperature and operating pressure.

Schmidt et al (J. Catal., 191, 62–74 (2000)) describes an oxidativeethane oxidation over a Pt—Sn/α-Al₂O₃ catalyst at a total feed flow rate(ethane, hydrogen and oxygen reactive components, nitrogen diluent) of 4to 16 standard litres per minute (SLPM), corresponding to a superficialfeed velocity of ˜22 to ˜88 cm s⁻¹ at standard temperature and operatingpressure. A small fall in ethylene yield was reported on raising the gasflow to the higher figure.

Holmen et al, Studies in Surf. Sci. and Catal., 119, 641–646 (1998)disclose the use of Pt and Pt/Rh gauze catalysts for oxidative ethanedehydrogenation. Experiments were conducted at a total gas feed rate of2 standard litres per minute over a Pt/Rh gauze (which corresponds tosuperficial velocities up to ˜265 cm s⁻¹ at standard temperature andoperating pressure). Although they report that the formation of olefins(selectivity) is favoured by short contact times, they also note thatconversion was reduced at high velocities when compared with results at˜19 cm s⁻¹ unless more heat was applied to the reactor externally.

WO 00/14035 discloses a process for the partial oxidation of paraffinichydrocarbons to form olefins. The process is carried out in the presenceof hydrogen and the use of gas hourly space velocities of greater than50,000 h⁻¹ to generally less than 6,000,000 h⁻¹ is disclosed. In oneexample there is disclosed the partial oxidation of ethane in thepresence hydrogen and a ceramic supported Pt/Cu catalyst at gas feedrates of up to 42 standard litres per minute and at a pressure of 1.68bara. This corresponds to superficial feed velocities up to ˜164 cm s⁻¹at standard temperature and operating pressure.

U.S. Pat. No. 4,940,826 discloses a catalytic oxidative dehydrogenationprocess with a hydrocarbon stream consisting of ethane, propane orbutane or mixtures thereof over platinum supported on cordieritemonolith or over a bed of platinum on alumina spheres. The total feedflow rates range from 16.0 to 55.0 standard litres per minutecorresponding to superficial feed velocities of ˜45 to ˜180 cm s⁻¹ atstandard temperature and operating pressure.

U.S. Pat. No. 5,382,741 discloses a catalytic oxidative dehydrogenationprocess carried out at elevated pressures (10 barg) over platinum andpalladium supported on a foam monolith or on a bed of alumina spheres.The hydrocarbon feeds exemplified are propane and naphtha. The totalfeed flow rates range from 2.1 SLPM at 1 bara to 280 SLPM at 11 bara,corresponding to superficial feed velocities of ˜44 to ˜240 cm s⁻¹ atstandard temperature and operating pressure.

SUMMARY OF THE INVENTION

The use of higher superficial feed velocities provides the advantagesthat the auto-thermal cracking process may be carried out using areduced number of reactors and also at a reduced risk of flashback.

Thus, it would be desirable to operate an auto-thermal cracking processusing higher superficial feed velocities than have previously been usedbut without incurring significant deterioration in hydrocarbon feedconversion and olefin yield.

Accordingly, the present invention provides a process for the productionof an olefin, said process comprising:

-   -   partially combusting in a reaction zone a mixture of a        hydrocarbon and an oxygen-containing gas in the presence of a        catalyst which is capable of supporting combustion beyond the        fuel rich limit of flammability to produce the olefin, wherein        the superficial feed velocity of said mixture is at least 300 cm        s⁻¹ at standard temperature and operating pressure.

According to a second aspect of the present invention, there is provideda process for the production of an olefin, said process comprising:

-   -   partially combusting in a reaction zone a mixture of a        hydrocarbon and an oxygen-containing gas in the presence of a        catalyst which is capable of supporting combustion beyond the        fuel rich limit of flammability to produce the olefin, wherein        the superficial feed velocity of said mixture is at least 250 cm        s⁻¹ at standard temperature and operating pressure and wherein        the catalyst is supported on a catalyst support.

The superficial feed velocity of the hydrocarbon and oxygen-containinggas mixture may be any practical superficial feed velocity, but wherethe catalyst is an unsupported catalyst, is at least 300 cm s⁻¹ andwhere the catalyst is a supported catalyst, is at least 250 cm s⁻¹.

Preferably the superficial feed velocity of the hydrocarbon andoxygen-containing gas mixture is in the range 300 cm s⁻¹ to 5000 cm s⁻¹.More preferably the superficial feed velocity is in the range 500 to3000 cm s⁻¹, even more preferably, in the range 600 to 2000 cm s⁻¹, andmost preferably, in the range 600 to 1200 cm s⁻¹; for example, in therange 600 to 700 cm s⁻¹.

The hydrocarbon may be any hydrocarbon which can be converted to anolefin, preferably a mono-olefin, under the partial combustionconditions employed.

The process of the present invention may be used to convert both liquidand gaseous hydrocarbons into olefins. Suitable liquid hydrocarbonsinclude naphtha, gas oils, vacuum gas oils and mixtures thereof.Preferably, however, gaseous hydrocarbons such as ethane, propane,butane and mixtures thereof are employed. Suitably, the hydrocarbon is aparaffin-containing feed comprising hydrocarbons having at least twocarbon atoms.

The hydrocarbon feed is mixed with any suitable oxygen-containing gas.Suitably, the oxygen-containing gas is molecular oxygen, air, and/ormixtures thereof. The oxygen-containing gas may be mixed with an inertgas such as nitrogen or argon.

Additional feed components may be included, if so desired. Suitably,methane, hydrogen, carbon monoxide, carbon dioxide or steam may beco-fed into the reactant stream.

Any molar ratio of hydrocarbon to oxygen-containing gas is suitableprovided the desired olefin is produced in the process of the presentinvention. The preferred stoichiometric ratio of hydrocarbon tooxygen-containing gas is 5 to 16, preferably, 5 to 13.5 times,preferably, 6 to 10 times the stoichiometric ratio of hydrocarbon tooxygen-containing gas required for complete combustion of thehydrocarbon to carbon dioxide and water.

The hydrocarbon is passed over the catalyst at a gas hourly spacevelocity of greater than 10,000 h⁻¹, preferably above 20,000 h⁻¹ andmost preferably, greater than 100,000 h⁻¹. It will be understood,however, that the optimum gas hourly space velocity will depend upon thepressure and nature of the feed composition.

Additionally, the use of a high superficial feed velocity in combinationwith a high gas hourly space velocity provides the advantage that theamount of catalyst and/or size of reactor(s) required to carry out theauto-thermal cracking process is minimised. Suitably, therefore, wherethe superficial feed velocity is above 300 cm/s, the gas hourly spacevelocity is preferably above 200,000/h.

In a preferred embodiment of the present invention, hydrogen is co-fedwith the hydrocarbon and oxygen-containing gas into the reaction zone.The molar ratio of hydrogen to oxygen-containing gas can vary over anyoperable range provided that the desired olefin product is produced.Suitably, the molar ratio of hydrogen to oxygen-containing gas is in therange 0.2 to 4, preferably, in the range 1 to 3.

Advantageously, it has been found that the use of a hydrogen co-feedallows, for a given feed throughput, the use of higher superficial feedvelocities than when the process is carried out in the absence ofhydrogen.

Hydrogen co-feeds are also advantageous because, in the presence of thecatalyst, the hydrogen combusts preferentially relative to thehydrocarbon, thereby increasing the olefin selectivity of the overallprocess.

Preferably, the reactant mixture of hydrocarbon and oxygen-containinggas (and optionally hydrogen co-feed) is preheated prior to contact withthe catalyst. Generally, the reactant mixture is preheated totemperatures below the autoignition temperature of the reactant mixture.

Advantageously, a heat exchanger may be employed to preheat the reactantmixture prior to contact with the catalyst. The use of a heat exchangermay allow the reactant mixture to be heated to high preheat temperaturessuch as temperatures at or above the autoignition temperature of thereactant mixture. The use of high pre-heat temperatures is beneficial inthat less oxygen reactant is required which leads to economic savings.Additionally, the use of high preheat temperatures can result inimproved selectivity to olefin product. It has also be found that theuse of high preheat temperatures enhances the stability of the reactionwithin the catalyst thereby leading to higher sustainable superficialfeed velocities.

It should be understood that the autoignition temperature of a reactantmixture is dependent on pressure as well as the feed composition: it isnot an absolute value. Typically, in auto-thermal cracking processes,where the hydrocarbon is ethane at a pressure of 2 atmospheres, apreheat temperature of up to 450° C. may be used.

The catalyst is any catalyst which is capable of supporting combustionbeyond the fuel rich limit of flammability. Suitably, the catalyst maybe a Group VIII metal. Suitable Group VIII metals include platinum,palladium, ruthenium, rhodium, osmium and iridium. Preferably, the GroupVIII metal is rhodium, platinum, palladium or mixtures thereof.Especially preferred are platinum, palladium or mixtures thereof.Typical Group VIII metal loadings range from 0.01 to 100 wt %,preferably, from 0.1 to 20 wt %, and more preferably, from 0.5 to 10 wt%, for example 1–5 wt %, such as 3–5 wt %.

Where a Group VIII metal catalyst is employed, it is preferably employedin combination with at least one promoter. The promoter may be selectedfrom elements of Groups IIIA, IVA and VA of the Periodic Table andmixtures thereof. Alternatively, the promoter may be a transition metal,which is different to the Group VIII metal(s) employed as the catalyticcomponent.

Preferred Group IIIA metals include Al, Ga, In and Tl. Of these, Ga andIn are preferred. Preferred Group IVA metals include Ge, Sn and Pb. Ofthese, Ge and Sn are preferred. The preferred Group VA metal is Sb. Theatomic ratio of Group VIII metal to the Group IIIA, IVA or VA metal maybe 1:0.1–50.0, preferably, 1:0.1–12.0, such as 1:0.3–5.

Suitable transition metal promoters may be selected from any one or moreof Groups IB to VIII of the Periodic Table. In particular, transitionmetals selected from Groups IB, IIB, VIB, VIIB and VIIIB of the PeriodicTable are preferred. Examples of such metals include Cr, Mo, W, Fe, Ru,Os, Co, Rh, Ir, Ni, Pt, Cu, Ag, Au, Zn, Cd and Hg. Preferred transitionmetal promoters are Mo, Rh, Ru, Ir, Pt. Cu and Zn. The atomic ratio ofthe Group VIII metal to the transition metal promoter may be 1:0.1–50.0,preferably, 1:0.1–12.0.

In one embodiment of the present invention, the catalyst comprises asingle promoter metal selected from Group IIIA, Group IVA, Group VB andthe transition metal series. For example, the catalyst may comprise asthe catalytic component, rhodium, platinum or palladium and as apromoter a metal selected from the group consisting of Ga, In, Sn, Ge,Ag, Au or Cu. Preferred examples of such catalysts include Pt/Ga, Pt/In,Pt/Sn, Pt/Ge, Pt/Cu, Pd/Sn, Pd/Ge, Pd/Cu and Rh/Sn. Of these Pt/Cu andPt/Sn are most preferred.

Where promoted Rh, Pd or Pt catalysts are employed, the Rh, Pt or Pd mayform between 0.01 and 5.0 wt %, preferably, between 0.01 and 2.0 wt %,and more preferably, between 0.05 and 1.5 wt % of the total weight ofthe catalyst. The atomic ratio of Rh, Pt or Pd to the Group IIIA, IVA ortransition metal promoter may be 1:0.1–50.0, preferably, 1:0.1–12.0. Forexample, atomic ratios of Rh, Pt or Pd to Sn may be 1:0.1 to 50,preferably, 1:0.1–12.0, more preferably, 1:0.2–4.0 and most preferably,1:0.5–2.0. Atomic ratios of Pt or Pd to Ge, on the other hand, may be1:0.1 to 50, preferably, 1:0.1–12.0, and more preferably, 1:0.5–8.0.Atomic ratios of Pt or Pd to Cu may be 1:0.1–3.0, preferably, 1:0.2–2.0,and more preferably, 1:0.3–1.5.

In another embodiment of the present invention, the promoter comprisesat least two metals selected from Group IIIA, Group IVA and thetransition metal series. For example, where the catalyst comprisesplatinum, it may be promoted with two metals from the transition metalseries, for example, palladium and copper. Such Pt/Pd/Cu catalysts maycomprise palladium in an amount of 0.01 to 5 wt %, preferably, 0.01 to 2wt %, and more preferably, 0.01 to 1 wt %. The atomic ratio of Pt to Pdmay be 1:0.1–10.0, preferably, 1:0.5–8.0, and more preferably,1:1.0–5.0. The atomic ratio of platinum to copper is preferably1:0.1–3.0, more preferably, 1:0.2–2.0, and most preferably, 1:0.5–1.5.

Alternatively, where the catalyst comprises platinum, it may be promotedwith one transition metal, and another metal selected from Group IIIA orGroup IVA of the periodic table. In such catalysts, palladium may bepresent in an amount of 0.01 to 5 wt %, preferably, 0.01 to 2.0 wt %,and more preferably, 0.05–1.0 wt % based on the total weight of suchcatalysts. The atomic ratio of Pt to Pd may be 1:0.1–10.0, preferably,1:0.5–8.0, and more preferably, 1:1.0–5.0. The atomic ratio of Pt to theGroup IIIA or IVA metal may be 1:0.1–60, preferably, 1:0.1–50.0.Preferably, the Group IIIA or IVA metal is Sn or Ge, most preferably,Sn.

For the avoidance of doubt, the Group VIII metal and the promoter in thecatalyst may be present in any form, for example, as a metal, or in theform of a metal compound, such as an oxide.

It should be understood that the actual concentrations of metal in thecatalysts tend not to be identical to the nominal concentrationsemployed in the preparation of the catalyst because not all of the metalemployed during the preparation of the catalyst becomes incorporatedinto the final catalyst composition. Thus, the nominal metalconcentrations may have to be varied to ensure that the desired actualmetal concentrations are achieved.

The catalyst employed in the present invention may be unsupported. Forexample, the catalyst may be in the form of a metal gauze. Preferably,however, the catalyst employed in the process of the present inventionmay be a supported catalyst. Although a range of support materials maybe used, ceramic supports are generally preferred. However, metalsupports may also be used.

Suitably, the ceramic support may be any oxide or combination of oxidesthat is stable at high temperatures of, for example, between 600° C. and1200° C. The ceramic support material preferably has a low thermalexpansion co-efficient, and is resistant to phase separation at hightemperatures.

Suitable ceramic supports include cordierite, lithium aluminium silicate(LAS), alumina (alpha—Al₂O₃), yttria stabilised zirconia, aluminiumtitanate, niascon, and calcium zirconyl phosphate.

The ceramic support may be wash-coated, for example, with gamma-Al₂O₃.

The structure of the support material is important, as the structure mayaffect flow patterns through the catalyst. Such flow patterns mayinfluence the transport of reactants and products to and from thecatalyst surface, thereby affecting the activity of the catalyst.Typically, the support material may be in the form of particles, such asspheres or other granular shapes or it may be in the form of a foam orfibre such as a fibrous pad or mat. Preferably, the form of the supportis a monolith which is a continuous multi-channel ceramic structure.Such monoliths include honeycomb structures, foams, or fibrous pads. Thepores of foam monolith structures tend to provide tortuous paths forreactants and products. Such foam monolith supports may have 20 to 80,preferably, 30 to 50 pores per inch. Channel monoliths generally havestraighter, channel-like pores. These pores are generally smaller, andthere may be 80 or more pores per linear inch of catalyst.

Preferred ceramic foams include lithium aluminium silicate.

Alternatively, the support may be present as a thin layer or wash coaton another substrate.

Preferred supports include gamma-alumina wash-coated lithium aluminiumsilicate foam and alumina spheres.

The catalyst employed in the present invention may be prepared by anymethod known in the art. For example, gel methods and wet-impregnationtechniques may be employed. Typically, the support is impregnated withone or more solutions comprising the metals, dried and then calcined inair. The support may be impregnated in one or more steps. Preferably,multiple impregnation steps are employed. The support is preferablydried and calcined between each impregnation, and then subjected to afinal calcination, preferably, in air. The calcined support may then bereduced, for example, by heat treatment in a hydrogen atmosphere.

The catalyst may be in the form of a fluidised or fixed bed. Preferably,a fixed bed catalyst is employed.

The process of the present invention may suitably be carried out at acatalyst exit temperature in the range 600° C. to 1200° C., preferably,in the range 850° C. to 1050° C. and, most preferably, in the range 900°C. to 1000° C.

The process of the present invention may be carried out at atmosphericor elevated pressure. Suitably, the pressure may be in the range from 0to 2 bara, preferably 1.5 to 2 bara, for example 1.8 bara. Elevatedpressures of, for example, 2 to 50 bara, may also be suitable.

Where the process of the present invention is carried out at elevatedpressure, the reaction products may be quenched as they emerge from thereaction chamber to avoid further reactions taking place.

Any coke produced in the process of the present invention may be removedby mechanical means, or by using one of the decoking methods such asthat described in EP-A-0 709 446, the contents of which are herebyincorporated by reference.

The degree of conversion of the hydrocarbon in the process of thepresent invention may be influenced by such factors as the nature of thefeed composition, the process conditions, the catalyst composition, thereactor and, in particular, heat losses from the reactor. High heatlosses can lead to lower hydrocarbon conversion as some of the energygenerated by the exothermic combustion reaction is lost to thesurroundings rather than being utilized to convert the hydrocarbon toolefin. External heating around the reaction zone can be employed tominimise heat losses and approach adiabatic operation. In the process ofthe present invention, the conversion of the hydrocarbon is generally atleast 30 mole %, preferably, at least 50 mole %, more preferably, atleast 60 mole %, such as at least 70 mole %.

The selectivity to olefin in the process of the present invention mayvary depending on such factors as the nature of the feed composition,the process conditions, the composition of the catalyst and the reactor.In the process of the present invention, selectivity to olefin istypically at least 60 g per 100 g hydrocarbon converted, preferably, atleast 70 g per 100 g hydrocarbon converted.

The invention will now be illustrated by way of example only and withreference to FIG. 1 and to the following examples.

BRIEF DESCRIPTION OF THE DRAWING

FIG. 1 represents in schematic form, apparatus suitable for use in theprocess of the present invention.

DETAILED DESCRIPTION OF THE INVENTION

FIG. 1 depicts an apparatus 10 comprising a quartz reactor 12 surroundedby an electrically-heated furnace 14. The reactor 12 is coupled to anoxygen-containing gas supply 16 and a hydrocarbon feed supply 18.Optionally, the hydrocarbon feed may comprise a co-feed such as hydrogenand a diluent such as nitrogen. In use, a catalyst which is capable ofsupporting combustion beyond the fuel rich limit of flammability 20 islocated within the reactor 12. The catalyst 20 is placed between a pairof ceramic foam heat shields 22, 24.

The furnace 14 is set to minimise heat losses, and the reactants areintroduced into the reactor via line 26. In use, as the reactantscontact the catalyst 20, some of the hydrocarbon feed combusts toproduce water and carbon oxides. The optional hydrogen co-feed alsocombusts to produce water. Both of these combustion reactions areexothermic, and the heat produced therefrom is used to drive thecracking of the hydrocarbon to produce olefin.

Catalyst Preparation Experiments

Experiment 1—Preparation of Catalyst A (0.7 wt % Pt)

The catalyst was prepared by impregnating a lithium aluminium silicatefoam support (30 pores per inch, ex Vesuvius Hi-Tech Ceramics Inc)having a high porosity alumina (HPA) wash-coat in a solution of(NH₃)₄Pt^(II)Cl₂. The (NH₃)₄Pt^(II)Cl₂ solution was prepared withsufficient salt to achieve a nominal Pt loading of 0.7 wt %. Thequantity of salt employed was that which would achieve the final targetloading if 100% of the platinum metal in the salt was taken up by thesupport material. The (NH₃)₄Pt^(II)Cl₂ was dissolved in a volume ofde-ionised water equivalent to three times the bulk volume of thesupport material. The support was impregnated with the platinumsolution, dried in air at 120° C. for ca. 30 minutes, then calcined inair at 450° C. for a further 30 minutes. The support was then allowed tocool to room temperature and the impregnation-drying-calcination cyclewas repeated until all of the platinum solution had been absorbed on tothe support (3–4 cycles were required.). The catalyst was then calcinedin air at 1200° C. for 6 hours (the temperature being increased from450° C. to 1200° C. at 2° C./min).

Experiment 2—Preparation of Catalyst B (3 wt % Pt)

The procedure of Experiment 1 was repeated using a (NH₃)₄Pt^(II)Cl₂solution of sufficient concentration to achieve a nominal Pt loading of3 wt %.

Experiment 3—Preparation of Catalyst C (5 wt % Pt)

The procedure of Experiment 1 was repeated using a (NH₃)₄Pt^(II)Cl₂solution of sufficient concentration to achieve a nominal Pt loading of5 wt %.

Experiment 4—Preparation of Catalyst D (3 wt % Pt, 1 wt % Cu)

The catalyst was prepared by impregnating a lithium aluminium silicatefoam support (30 pores per inch, ex Vesuvius Hi-Tech Ceramics Inc)having an HPA wash-coat with a solution of 1) (NH₃)₄Pt^(II)Cl₂, and 2)Cu(NO₃)₂. Prior to the impregnation process the LAS-HPA foam support wascalcined in air at 1200° C.

Solutions of (NH₃)₄Pt^(II)Cl₂, and Cu(NO₃)₂ in de-ionised water wereprepared with sufficient salt to achieve nominal Pt and Cu loadings of 3wt % and 1 wt %, respectively. The quantity of salt dissolved wasequivalent to that needed to achieve the final target platinum andcopper loading if 100% of the platinum and copper were to be recoveredon the final catalyst. The volumes of de-ionised water used for thesolutions were equal to three times the bulk volume of the supportmaterial.

The support was alternately impregnated with the platinum- andcopper-containing solutions. Between each impregnation the support wasdried in air at 120° C. for ca. 30 minutes, calcined in air at 450° C.for a further 30 minutes, then cooled to room temperature for thesubsequent impregnation. The impregnation-drying-calcination cycles wererepeated until all the impregnation solutions had been absorbed onto thesupport.

The impregnated support was then dried, and then finally calcined in airfor 6 hours at 600° C. Immediately prior to use in the auto-thermalcracking reaction the catalyst was reduced in-situ using ca. 2 nl/min ofhydrogen and 2 nl/min of nitrogen. The reduction temperature wasmaintained for 1 hour at 750° C.

Experiment 5—Preparation of Catalyst E (3 wt % Pt, 1 wt % Cu)

The procedure of Experiment 4 was repeated, except that the finalcalcination in air was carried out at 1200° C.

Experiment 6—Preparation of Catalyst F (3 wt % Pt, 1 wt % Cu)

The procedure of Experiment 4 was repeated, except that the finalcalcinations in air was carried out at 1200° C., and the reduction stepwas omitted.

Experiment 7—Preparation of Catalyst G (2 wt % Pt, 4 wt % Sn)

The procedure of Experiment 4 was repeated using a (NH₃)₄Pt^(II)Cl₂solution of sufficient concentration to give a nominal Pt loading of 2wt % and a SnCl₂/dil HCl solution of sufficient concentration to give anominal Sn loading of 4 wt %.

Experiment 8—Preparation of Catalyst H (4 wt % Pt, 4 wt % Sn)

The procedure of Experiment 4 was repeated using a (NH₃)₄Pt^(II)Cl₂solution of sufficient concentration to give a nominal Pt loading of 4wt % and a SnCl₂/dil HCl solution of sufficient concentration to give anominal Sn loading of 4 wt %.

EXAMPLE 1 Auto-thermal Cracking of Ethane in the Presence of Hydrogen atAtmospheric Pressure

The catalysts A to H as prepared in Experiments 1 to 8 above and havingdimensions 15 mm diameter by 30 mm depth, a porosity of 30 pores perinch and a volume of 5.30 cm³ were placed in an apparatus as describedfor FIG. 1. The reactor had an internal diameter of 15 mm. Oxygen,ethane, hydrogen and nitrogen as diluent (10 vol %) were contacted withthe catalyst under the conditions shown in Table 1 below. The ratio ofhydrogen to oxygen was 2:1 (v/v); the oxygen: ethane feed ratio was 0.65(wt/wt) (1.00:2.04 v/v). The reaction was carried out at atmosphericpressure.

The product composition was analysed by gas chromatography fitted withthermal conductivity and flame ionization detectors. Gas feed rates werecontrolled by thermal mass flow controllers (ex Bronkhorst HiTec).

The electrically-heated furnace surrounding the reactor and catalyst wasset to 850° C. to minimise heat losses from the catalyst/reaction zone.

From analysis of the feed and product flow rates and compositions thefollowing parameters were calculated:

Conversion

-   Ethane conversion %=ethane feed (g/min)−ethane in effluent    (g/min)/ethane feed (g/min)*100-   Oxygen conversion %=oxygen feed (g/min)−oxygen in effluent    (g/min)/oxygen feed (g/min)*100-   Change in ethane conversion (%/(cm/s))=(ethane conversion @ higher    velocity)−(ethane conversion @ lower velocity)/higher velocity    (cm/s)−lower velocity (cm/s)-   Change in oxygen conversion (%/(cm/s))=(oxygen conversion @ higher    velocity)−(oxygen conversion @ lower velocity)/higher velocity    (cm/s)−lower velocity (cm/s)    Selectivity

$\begin{matrix}\text{Ethylene selectivity} \\\text{(g per 100 g ethane converted)}\end{matrix} = {100 \times \frac{\text{ethylene in product (g/min)}}{\text{ethane in feed (g/min)} - \text{ethane in product (g/min)}}}$Yeild

$\begin{matrix}\text{Ethylene yield} \\\text{(g per 100 g ethane feed)}\end{matrix} = {100 \times \frac{\text{ethylene in product (g/min)}}{\text{ethane in feed (g/min)}}}$

The results are given in Table 1.

Table 1 clearly shows that the decrease in conversion to olefin onincreasing the superficial feed velocity from approximately 200 to 670cm s⁻¹ is relatively small.

TABLE 1 Catalyst A Catalyst D Catalyst E (0.7 wt % Catalyst B Catalyst C(3 wt % Pt, (3 wt % Pt, Pt) (3 wt% Pt) (5 wt% Pt) 1 wt % Cu) 1 wt % Cu)feed 195 119 208 133 186 132 222 136 232 119 temperature ° C. cat face546 233 663 408 682 631 1034 746 1033 778 temp ° C. cat exit 940 795 921837 903 860 955 966 950 901 temp ° C. adiabatic 790 664 799 690 783 667811 749 814 748 temp ° C. cat face −313 −255 −251 −288 −255 temp ° C.change cat exit −145 −84 −43 11 −49 temp ° C. change heat loss % 13.0831.03 10.28 24.83 11.09 24.25 10.29 19.55 9.44 17.07 feed rates Totalfeed 22.03 71.67 22.20 71.23 22.30 70.89 22.18 70.82 22.18 70.76 ratenl/min GSHV /h 249329 811138 251253 806159 252384 802311 251026 801518251026 800839 Superficial 208 676 209 672 210 669 209 668 209 667 gascm/s velocity at standard temperature and operating pressure conversionsethane % 76.08 35.5 78.39 44.24 74.25 43.12 81.24 63.45 82.23 62.96oxygen % 98.3 84.93 96.53 83.19 95.98 81.32 98.31 94.85 97.86 92.86change in conversion ethane % per 8.67 7.39 6.79 3.88 4.21 cm/s oxygen %per 2.86 2.88 3.20 0.75 1.09 cm/s Ethylene yield 52.60 24.98 53.73 31.7851.27 31.66 56.64 47.34 56.84 47.08 (g/100 g ethane feed) Ethylene(g/100 g 69.13 70.37 68.55 71.84 69.05 73.42 69.73 74.60 69.12 74.78selectivity ethane converted) Catalyst F Catalyst G Catalyst H (3 wt%Pt, (2 wt % Pt, (4 wt % Pt, 1 wt % Cu) 4 wt % Sn) 4 wt % Sn) feed 186121 161 138 208 127 temperature ° C. cat face 971 621 901 350 1032 690temp ° C. cat exit 950 930 937 969 967 958 temp ° C. adiabatic 809 742807 776 810 741 temp ° C. cat face −350 −551 −342 temp ° C. change catexit −20 32 −9 temp ° C. change heat loss % 7.96 19.36 9.75 11.52 9.9418.55 feed rates Total feed 22.16 71.15 17.72 69.41 22.15 70.84 ratenl/min GSHV /h 250800 805253 200549 785561 250687 801745 Superficial 209671 167 655 209 668 gas cm/s velocity at standard temperature andoperating pressure conversions ethane % 80.77 60.94 79.66 71.31 80.7660.59 oxygen % 97.81 93.23 98.88 94.52 98.08 91.5 change in conversionethane % per 4.29 1.71 4.39 cm/s oxygen % per 0.99 0.89 1.43 cm/sEthylene yield 56.49 45.65 57.15 52.16 56.14 45.24 (g/100 g ethane feed)Ethylene (g/100 g 69.94 74.91 71.74 73.15 69.52 74.66 selectivity ethaneconverted)Experiment 9—Preparation of Catalyst I (3 wt % Pt)

The catalyst was prepared by impregnating alumina spheres (1.8 mmdiameter, ex Condea) with a solution of (NH3)4PtIICl2. Prior toimpregnation the spheres were calcined in air to 1200° C. for 6 hours toremove any residual porosity.

A solution of (NH3)4PtIICl2 in de-ionised water was prepared withsufficient salt to achieve a nominal Pt loading of 3 wt %. The quantityof salt dissolved was equivalent to that needed to achieve the finaltarget platinum loading if 100% of the platinum were to be recovered onthe final catalyst. The volume of de-ionised water used for the solutionwas equal to the bulk volume of the support material.

The support was impregnated with the platinum solution, dried in air at120° C. for ca. 30 minutes, then calcined in air at 450° C. for afurther 30 minutes and then cooled to room temperature. Theimpregnation-drying-calcination cycle was repeated until all of theplatinum solution had been absorbed on to the support (1–2 cycles wererequired.). After the final calcination at 450° C., the catalyst wasfurther calcined in air at 1200° C. for 6 hours (the temperature beingincreased from 450° C. to 1200° C. at 5° C./min) and then allowed tocool to room temperature.

Experiment 10—Preparation of Catalyst J (3 wt % Pt, 1 wt % Cu)

The catalyst was prepared by impregnating alumina spheres (1.8 mmdiameter, ex Condea) with solutions of I) (NH3)4PtIICl2 and 2) Cu(NO₃)₂.Prior to impregnation the spheres were calcined in air to 1200° C. for 6hours to remove any residual porosity.

Solutions of (NH₃)₄Pt^(II)Cl₂, and Cu(NO₃)₂ in de-ionised water wereprepared with sufficient salt to achieve nominal Pt and Cu loadings of 3w % and 1 wt %, respectively. The quantity of salt dissolved wasequivalent to that needed to achieve the final target platinum andcopper loading if 100% of the platinum and copper were to be recoveredon the final catalyst. The volumes of de-ionised water used for thesolutions were equal to the bulk volume of the support material.

The support was alternately impregnated with the platinum and coppersolutions. Between each impregnation the support was dried in air at120° C. for ca. 30 minutes, calcined in air at 450° C. for a further 30minutes, then cooled to room temperature for the subsequentimpregnation. The impregnation-drying-calcination cycles were repeateduntil all of the impregnation solutions had been absorbed onto thesupport.

After the final calcination at 450° C., the catalyst was furthercalcined in air at 600° C. for 6 hours (the temperature being increasedfrom 450° C. to 600° C. at 5° C./min) and then allowed to cool to roomtemperature.

Prior to use in the auto-thermal cracking reaction the catalyst wasreduced using ca. 2 nl/min of hydrogen and 2 nl/min of nitrogen and at atemperature of 750° C. This reduction temperature was maintained for 1hour after which the catalyst was allowed to cool to room temperatureunder nitrogen and then transferred to the reactor.

Experiment 11—Preparation of Catalyst K (1 wt % Pt, 4 wt % Sn)

The procedure of Experiment 10 was repeated using a (NH₃)₄Pt^(II)Cl₂solution of sufficient concentration to give a nominal Pt loading of 1wt % and a SnCl₂/dil HCl solution of sufficient concentration to give anominal Sn loading of 4 wt %.

Experiment 12—Preparation of Catalyst L (3 wt % Pt)

The catalyst was prepared by the procedure of Experiment 9 except that alithium aluminium silicate foam support (30 pores per inch, ex VesuviusHi-Tech Ceramics Inc) was used in place of the alumina and the(NH₃)₄Pt^(II)Cl₂ was dissolved in a volume of deionised water equivalentto three times the bulk volume of the support material.

Experiment 13—Preparation of Catalyst M (3 wt % Pt, 1 wt % Cu)

The procedure of Experiment 10 was repeated using a (NH3)4PtIICl2solution of sufficient concentration to give a nominal loading of 3 wt %Pt and a Cu(NO₃)₂ solution of sufficient concentration to give a nominalloading of 1 wt % Cu. In addition the alumina spheres were replaced by alithium aluminium silicate foam (30 pores per inch, ex Vesuvius Hi-TechCeramics Inc) and the volumes of de-ionised water used for the platinumand copper solutions were equivalent to three times the bulk volume ofthe support material.

EXAMPLE 2 Auto-thermal Cracking of Ethane in the Presence of Hydrogen atElevated Pressure

The catalysts I to K as prepared in Experiments 9 to 11 above and havingdimensions 15 mm diameter by 60 mm depth, and a volume of 10.60 cm³ wereplaced in a metallic reactor (internal diameter 15 mm) with a quartzlining and fitted with a pressure jacket. Catalysts I to K were testedas packed beds of spheres supported on an alumina foam block ofdimensions 15 mm diameter, 10 mm depth and of porosity 30 pores perinch.

Catalyst L as prepared in Experiment 12 above was tested as a ceramicfoam bed within a metallic reactor fitted with a pressure jacket (beddiameter 18 mm, bed depth 60 mm, bed volume 15.27 cm³).

Catalyst M as prepared in Experiment 13 above was tested as a ceramicfoam bed in a quartz lined metallic reactor fitted with a pressurejacket (bed diameter 15 mm, bed depth 60 mm, bed volume 10.60 cm³).

The pressure jacket was not externally heated.

The catalysts (I-M) were heated to approximately 200° C. under nitrogenat reaction pressure. Oxygen, ethane, hydrogen and nitrogen as diluent(10 vol %) preheated to 180–200° C. were then contacted with thecatalyst under the conditions shown in Table 2 below. The ratio ofhydrogen to oxygen was 2:1 (v/v); the oxygen:ethane feed ratio forcatalysts I-K and M was 1.00:1.77 v/v; the oxygen:ethane feed ratio forcatalyst K was 1.00:2.34 v/v. The reaction pressures are shown in Table2.

The product composition was analysed by gas chromatography fitted withthermal conductivity and flame ionization detectors. Gas feed rates werecontrolled by thermal mass flow controllers (ex Bronrkhorst HiTec BV)

From analysis of the feed and product flow rates and compositions, theethane and oxygen conversions, ethylene selectivity and yield werecalculated using the equations given in Example 1.

The results are given in Table 2.

TABLE 2 Catalyst J Catalyst K Catalyst M Catalyst I (3 wt % Pt, 1 (1 wt%Pt, 4 Catalyst L (3 wt % Pt, 1 (3 wt % Pt) wt % Cu) wt % Sn) (3 wt % Pt)wt % Cu) preheat 186 191 191 194 178 195 217 257 164 186 temperature (°C.) cat face 925 900 1041 1045 1054 1045 298 300 340 265 temp (° C.) catexit 898 944 944 990 886 1019 911 897 850 969 temp (° C.) temp change−25 4 −9 2 −75 front temp change 46 46 133 −14 119 base adiabatic temp717 799 794 820 766 854 738 770 707 804 heat loss % 39.47 18.58 24.5016.10 33.74 10.14 15.50 16.14 42.03 23.16 pressure 1.3 1.3 1.3 1.44 1.31.3 1.8 1.8 1.3 1.3 (bara) feed rates total (nl/min) 40.00 109.74 94.68147.81 38.46 144.16 104.59 257.02 35.61 123.62 GSHV (/h) 226354 621001535779 836433 217639 815779 411013 1010025 201511 699546 superficial gas290 796 687 968 279 1046 381 935 258 897 velocity* (cm/s) conversionsethane (%) 56.14 79.98 77.78 84.7 69.44 92.68 65.7 67.56 52.99 80.75oxygen (%) 95.86 94.52 97.02 96.82 98.41 98.06 96.98 95.97 95.98 95.19ethylene 41.36 52.39 54.07 55.13 52.11 55.91 44.21 44.03 38.84 49.27yield (g per 100 g ethane feed) ethylene 73.68 65.84 69.51 65.09 75.0460.33 67.30 65.16 73.30 61.01 selectivity (g per 100 g ethane converted)*at standard temperature and operating pressureFrom Table 2 it is evident that the use of superficial feed velocitiesabove 250 cm/s with supported catalysts produces acceptable ethyleneconversions and yields. It can be seen that in these examples the heatloss at the lower superficial feed velocities is large (no externalheating to compensate for losses to the immediate environment of thecatalyst). As the superficial feed velocity is increased the heat lossesdecline and the losses to the environment become less significant as afraction of the enthalpy of the products. As a consequence ethaneconversion is seen to rise and ethylene yields are maintained.Experiments 14–15—Preparation of Catalysts N and P

Catalysts N and P were each prepared by the procedure of Experiment 10except that (i) solutions of (NH3)4PtIICl2, and (NH3)4PdIICl2 ofsufficient concentration to achieve nominal Pt and Pd loadings for eachcatalyst as given in Table 3 were used (ii) a lithium aluminium silicatefoam support (30 pores per inch, ex Vesuvius Hi-Tech Ceramics Inc) wasused in place of the alumina spheres, iii) the volumes of de-ionisedwater used for the platinum and palladium solutions were equal to threetimes the bulk volume of the support material and (iv) there was nohydrogen reduction treatment.

EXAMPLE 3 Auto-thermal Cracking of Ethane in the Absence of Hydrogen atAtmospheric Pressure

The catalysts N and P as prepared in Experiments 14 and 15 above andhaving dimensions and a volume as shown in Table 3 were placed in ametallic reactor (internal diameter 15 mm) with a quartz lining. Oxygen,ethane, and nitrogen were then contacted with the catalyst under theconditions shown in Table 3 below. The reaction was carried out atatmospheric pressure.

The product composition was analysed by gas chromatography fitted withthermal conductivity and flame ionization detectors. Gas feed rates werecontrolled by thermal mass flow controllers (ex Bronkhorst HiTec BV)

From analysis of the feed and product flow rates and compositions theethane and oxygen conversions, ethylene selectivity and yield werecalculated using the equations given in Example 1.

TABLE 3 Catalyst N Catalyst P Pt (wt %) 0.23 0.23 2.06 2.06 Pd (wt %)0.11 0.11 0.38 0.38 Catalyst volume (cm3) 5.30 5.30 5.30 5.30 Catalystdepth (mm) 30 30 30 30 Catalyst diameter (mm) 15 15 15 15 Preheattemperature (° C.) 150 150 150 150 Cat face temp (° C.) 562 409 544 531Cat exit temp (° C.) 853 886 828 996 Temp change front (° C.) −153 −13Temp change exit (° C.) 33 168 Adiabatic temp (° C.) 712 786 701 894Heat loss (%) 29.86 10.41 31.72 10.35 Ethane:oxygen (v/v) 1.93 1.96 1.931.59 Nitrogen:oxygen (v/v) 0.42 0.43 0.42 0.15 Total feed rate (nl/min)13.63 38.99 13.63 38.98 GHSV (/h) 154260 441277 154260 441163Superficial gas velocity* (cm/s) 129 368 129 368 Ethane conversion (%)67.68 83.28 61.02 98.61 Oxygen conversion (%) 98.62 98.54 98.8 99.66Ethylene yield (g per 100 g 39.26 47.84 37.44 35.71 ethane feed)Ethylene selectivity (g per 100 g 58.15 57.32 60.17 36.17 ethaneconverted) *at standard temperature and operating pressure

1. A process for the production of an olefin, said process comprising:partially combusting in a reaction zone a mixture of a hydrocarbon andan oxygen-containing gas in the presence of a catalyst which is capableof supporting combustion beyond the fuel rich limit of flammability toproduce the olefin, wherein the superficial feed velocity of saidmixture is at least 300 cm s⁻¹ at standard temperature and operatingpressure, and wherein the process is carried out at a pressure in therange 1.3 to 50 bara and the reaction zone is not externally heated. 2.A process according to claim 1 wherein the superficial feed velocity ofthe hydrocarbon and oxygen-containing gas mixture is in the range 300 to5000 cm/s.
 3. A process according to claim 1 wherein the hydrocarbon isa paraffin-containing hydrocarbon feed having at least two carbon atoms.4. A process according to claim 3 in which the hydrocarbon is selectedfrom the group consisting of ethane, propane, butane, naphtha, gas oil,vacuum gas oil and mixtures thereof.
 5. A process according to claim 1wherein the molar ratio of hydrocarbon to the oxygen-containing gas is 5to 16 times the stoichiometric ratio of hydrocarbon to oxygen-containinggas required for complete combustion to carbon dioxide and water.
 6. Aprocess according to claim 1 in which hydrogen is co-fed into thereaction zone.
 7. A process according to claim 6 in which the molarratio of hydrogen to oxygen-containing gas is in the range 0.2 to
 4. 8.A process according to claim 1 wherein the process is conducted at a gashourly space velocity of greater than 10,000/h.
 9. A process accordingto claim 1 in which the catalyst comprises a Group VIII metal.
 10. Aprocess according to claim 9 wherein the Group VIII metal is selectedfrom rhodium, platinum, palladium and mixtures thereof.
 11. A processaccording to claim 9 in which the Group VIII metal catalyst comprises atleast one promoter.
 12. A process according to claim 11 wherein the atleast one promoter is selected from the group consisting of elements ofGroups IIIA, IVA, VA of the Periodic Table and mixtures thereof and atransition metal which is a different metal to the Group VIII metalemployed as catalyst.
 13. A process according to claim 12 wherein thepromoter is selected from tin and copper.
 14. A process according toclaim 11 wherein the Group VIII metal is platinum and the promoter isselected from tin and copper.
 15. A process according to claim 11wherein the atomic ratio of Group VIII metal to promoter is in the range1:0.1–50.0.
 16. A process according to claim 1 wherein the catalyst issupported.
 17. A process according to claim 16 wherein the support is aceramic support.
 18. A process according to claim 17 wherein the supportis in the form of a monolith or particles.
 19. A process according toclaim 18 wherein the monolith is a foam or a fibre.
 20. A processaccording to claim 1 in which the hydrocarbon and oxygen-containing gasmixture is preheated to a temperature below the auto-ignitiontemperature of the mixture.